High-pressure polymerization process of ethylenically unsaturated monomers in a tubular reactor

ABSTRACT

The present technology relates to a process for polymerizing or copolymerizing ethylenically unsaturated monomers in the presence of free-radical polymerization initiators, wherein the polymerization is carried out in a continuously operated tubular reactor at temperatures from 100° C. to 350° C. and pressures from 180 MPa to 340 MPa, with a specific reactor surface area A sp  of 2 m 2 /(t/h) to 5.5 m 2 /(t/h), and the tubular reactor has a specific ratio RD sp  of 0.0050 MPa −1  to 0.0069 MPa −1  and an inner surface which has a surface roughness Ra of 2 μm or less.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a Continuation Application of co-pending U.S. patentapplication Ser. No. 15/240,084, filed Aug. 18, 2016, which is aContinuation Application of U.S. patent application Ser. No. 15/044,855,filed Feb. 16, 2016, now U.S. Pat. No. 9,441,057, issued Sep. 13, 2016,which claims the benefit of priority to European Patent Application No.15178042.6 filed on Jul. 23, 2015, which is incorporated herein byreference in its entirety.

FIELD OF THE INVENTION

The present disclosure provides a process for polymerizing orcopolymerizing ethylenically unsaturated monomers in the presence offree-radical polymerization initiators at temperatures from 100° C. to350° C. and pressures from 180 MPa to 340 MPa in a continuously operatedtubular reactor.

BACKGROUND OF THE INVENTION

Polyethylene is the most widely used commercial polymer and can beprepared by different processes. Polymerization in the presence offree-radical initiators at elevated pressures was the method firstdiscovered to obtain polyethylene and continues to be of commercialrelevance for the preparation of low density polyethylene (LDPE).

A common set-up of a plant for preparing low density polyethylenecomprises a polymerization reactor and reaction components that may bepressurized by a combination of two compressors, a primary compressorand a secondary compressor. At the end of the polymerization sequence, ahigh-pressure polymerization unit may further include apparatuses likeextruders and granulators for pelletizing the resulting polymer.Furthermore, such a polymerization unit generally may also comprisemeans for feeding monomers and comonomers, free-radical initiators,chain transfer agents and/or other substances at one or more positionsinto the polymerization reaction. A process and an apparatus for themanufacture of ethylene polymers and copolymers under high pressures aredisclosed in WO 2007/018871 A1.

A characteristic of the radically initiated polymerization ofethylenically unsaturated monomers under high pressure is that theconversion of the monomers is often incomplete, as only about 10% to 50%of the dosed monomers are converted. The resulting reaction mixture mayleave the reactor through a pressure control valve, often designated asa let-down valve, and may then be separated into polymeric and gaseouscomponents with the unreacted monomers being recycled. To avoidunnecessary decompression and compression steps, the separation intopolymeric and gaseous components is usually carried out in two stages.The monomer-polymer mixture leaving the reactor is transferred to afirst separating vessel, frequently called a high-pressure productseparator, in which the separation into polymeric and gaseous componentsis carried out at a pressure that allows recycling of the ethylene andcomonomers separated from the monomer-polymer mixture to the reactionmixture at a position between the primary compressor and the secondarycompressor. At the first separation vessel operating conditions, thepolymeric components within the separating vessel are in liquid state.The liquid phase obtained in the first separating vessel is transferredto a second separation vessel, frequently called a low-pressure productseparator, in which further separation into polymeric and gaseouscomponents takes place at lower pressure. The ethylene and comonomersseparated from the mixture in the second separation vessel are fed tothe primary compressor where they are compressed to the pressure of thefresh ethylene feed, combined with the fresh ethylene feed and thejoined streams are further pressurized to the pressure of thehigh-pressure gas recycle stream.

The recycling of unreacted monomers to the inlet of the reactor andconducting the recycle at high pressures, which may reduce the need forre-compressing, are measures that may improve the economics ofhigh-pressure polymerization processes. Nonetheless, carrying out therecycling requires considerable efforts commensurate with the amount ofmonomer converted to polymer per pass of the reactor. Consequently,there is a demand for conducting high-pressure polymerizations formaximizing the conversion of monomers per pass of the reactor to producethe targeted low density polyethylene grades. However, the possibilitiesfor influencing the conversion of monomers by changing polymerizationconditions are limited since the properties and the structure of theresulting ethylene homopolymers or copolymers, such as molecular weight,molecular weight distribution and the amount of short- and long-chainbranching, depend strongly on the reaction parameters.

Furthermore, various articles and applications of low densitypolyethylenes such as blown films often require a narrow molecularweight distribution of the low density polyethylene for achieving a goodbalance of optical and mechanical properties. Accordingly, there is ademand for low density polyethylenes with narrow molecular weightdistributions in a high-pressure polymerization process.

Hence, there is a need to overcome the disadvantages of the prior artand to provide a process which makes it possible to polymerize orcopolymerize ethylenically unsaturated monomers in a tubular reactorwith a high conversion of monomers to polymer per pass of the reactor.Furthermore, the process should allow for the preparation of low densitypolyethylenes with a high conversion rate per pass of the reactorwithout introducing detrimental effects in the article forming processesor on properties of the produced low density polyethylenes.

SUMMARY OF THE INVENTION

The present disclosure provides for a process for polymerizing orcopolymerizing ethylenically unsaturated monomers in the presence offree-radical polymerization initiators, wherein the polymerization iscarried out at temperatures from 100° C. to 350° C. and pressures in therange from 180 MPa to 340 MPa in a continuously operated tubularreactor, wherein the reaction gas composition is brought to thepolymerization pressure by a combination of a primary compressor and asecondary compressor and the compressed reaction gas composition is fedto the inlet of the tubular reactor or the reaction gas composition issplit in a main stream, which is fed to the inlet of the tubularreactor, and one or more side streams, which are fed to the tubularreactor downstream of the inlet of the tubular reactor, and wherein thepolymerization is conducted with a specific reactor surface area A_(sp)of 2 m²/(t/h) to 5.5 m²/(t/h), the specific reactor surface area A_(sp)being the ratio of the area of the inner surface of the tubularpolymerization reactor to the feed rate of reaction gas composition tothe tubular reactor, and wherein the tubular reactor has a specificratio RD_(sp) of from 0.0050 MPa⁻¹ to 0.0069 MPa⁻¹, RD_(sp) being theratio of the outer diameter d_(o) to the inner diameter d_(i) divided bythe design pressure p_(des) according to equation

${{RD}_{sp} = \frac{d_{o}}{d_{i}*p_{des}}},$

and the tubular reactor has an inner surface which has a surfaceroughness Ra of 2 μm or less, determined according to DIN EN ISO4287:2010.

In some embodiments, the tubular reactor has an inner diameter d_(i) offrom 50 to 120 mm.

In some embodiments, the design pressure p_(des) of the tubular reactoris from 240 MPa to 400 MPa.

In some embodiments, the feed rate of the reaction gas composition tothe tubular reactor is 80 t/h to 210 t/h.

In some embodiments, the tubular reactor is composed of tubes of alength of 5 m to 25 m.

In some embodiments, the entire reaction gas composition provided by thesecondary compressor is fed to the inlet of the tubular reactor.

In some embodiments, 30-90% by weight of the reaction gas compositionprovided by the secondary compressor is fed to the inlet of the tubularreactor and 10-70% by weight of the reaction gas composition provided bythe secondary compressor is fed as one or more side streams to thetubular reactor downstream of the inlet of the tubular reactor.

In some embodiments, the polymerization is carried out in the presenceof a chain transfer agent.

In some embodiments, the chain transfer agent comprises at least onealdehyde or at least one ketone.

In some embodiments, the amount of added aldehydes and ketones is from0.4 kg/t of prepared polymer to 10 kg/t of prepared polymer.

In some embodiments, propionic aldehyde is used as the sole chaintransfer agent or the chain transfer agent is a mixture of propionicaldehyde and one or more olefinic hydrocarbons.

In some embodiments, the present disclosure provides a process forpreparing a shaped article, wherein a polymer prepared by a process asdescribed above is converted into the shaped article.

In some embodiments, the shaped article is a film.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows schematically a set-up for a polymerization reactor inwhich the entire reaction gas composition compressed by the secondarycompressor is fed as one stream to the inlet of the tubular reactor.

FIG. 2 shows schematically a set-up for a polymerization reactor inwhich a part of the reaction gas composition compressed by the secondarycompressor is fed to the inlet of the tubular reactor and a part is fedas side streams to the tubular reactor downstream of the inlet of thetubular reactor.

DETAILED DESCRIPTION OF THE INVENTION

The present disclosure refers to a process for polymerizing orcopolymerizing ethylenically unsaturated monomers in the presence offree-radical polymerization initiators in a continuously operatedtubular reactor at temperatures from 100° C. to 350° C. and pressuresfrom 180 MPa to 340 MP. The present disclosure describes a process whichgives a high conversion of monomers to polymer per pass of the reactor.

The high-pressure polymerization may be a homopolymerization of ethyleneor a copolymerization of ethylene with one or more other monomers,provided that these monomers are free-radically copolymerizable withethylene under high pressure. Examples of suitable copolymerizablemonomers are α,β-unsaturated C₃-C₈-carboxylic acids, in particularmaleic acid, fumaric acid, itaconic acid, acrylic acid, methacrylic acidand crotonic acid, derivatives of β,β-unsaturated C₃-C₈-carboxylicacids, e.g. unsaturated C₃-C₁₅-carboxylic esters including esters ofC₁-C₆-alkanols, or anhydrides such as methyl methacrylate, ethylmethacrylate, n-butyl methacrylate or tert-butyl methacrylate, methylacrylate, ethyl acrylate, n-butyl acrylate, 2-ethylhexyl acrylate,tert-butyl acrylate, methacrylic anhydride, maleic anhydride or itaconicanhydride, and 1-olefins such as propene, 1-butene, 1-pentene, 1-hexene,1-octene or 1-decene. In addition, vinyl carboxylates such as vinylacetate can be used as comonomers.

In the case of copolymerization, the proportion of comonomer orcomonomers in the reaction mixture is from 1 to 50% by weight, such asfrom 3 to 40% by weight, based on the amount of monomers, i.e. the sumof ethylene and other monomers. Depending on the type of comonomer, thecomonomers may be fed at more than one point into the reactor set-up.For instance, the comonomers may be fed to the suction side of thesecondary compressor.

For the purposes of the present disclosure, polymers or polymericmaterials are all substances which are made up of at least two monomerunits. The polymers or polymeric materials are preferably low densitypolyethylenes having an average molecular weight M_(n) of more than 20000 g/mole. The term low density polyethylene is meant to includeethylene homopolymers and ethylene copolymers. The present disclosurecan also be advantageously employed in the preparation of oligomers,waxes and polymers having a molecular weight M_(n) of less than 20 000g/mole.

Possible initiators for starting the free-radical polymerization in therespective reaction zones are all substances that can produce radicalspecies under the conditions in the polymerization reactor, for example,oxygen, air, azo compounds or peroxidic polymerization initiators. Insome embodiments of the disclosure the polymerizations are carried outby using oxygen, either fed in the form of pure O₂ or as air. In suchcases, the initiator may first be mixed with the ethylene feed and thenfed to the reactor such that it is not only possible to feed a streamcomprising monomer and oxygen to the beginning of the polymerizationreactor but also to one or more points along the reactor, therebycreating two or more reaction zones. Initiation using organic peroxidesor azo compounds also represents embodiments of the present disclosure.Examples of suitable organic peroxides are peroxy esters, peroxy ketals,peroxy ketones and peroxycarbonates, e.g. di(2-ethylhexyl)peroxydicarbonate, dicyclohexyl peroxydicarbonate, diacetylperoxydicarbonate, tert-butyl peroxyisopropylcarbonate, di-sec-butylperoxydicarbonate, di-tert-butyl peroxide, di-tert-amyl peroxide,dicumyl peroxide, 2,5-dimethyl-2,5-di-tert-butylperoxyhexane, tert-butylcumyl peroxide, 2,5-dimethyl-2,5-di(tert-butylperoxy)hex-3-yne,1,3-diisopropyl monohydroperoxide or tert-butyl hydroperoxide,didecanoyl peroxide, 2,5-dimethyl-2,5-di(2-ethylhexanoylperoxy)hexane,tert-amyl peroxy-2-ethylhexanoate, dibenzoyl peroxide, tert-butylperoxy-2-ethylhexanoate, tert-butyl peroxydiethylacetate, tert-butylperoxydiethylisobutyrate, tert-butyl peroxy-3,5,5-trimethylhexanoate,1,1-di(tert-butylperoxy)-3,3,5-trimethylcyclohexane,1,1-di(tert-butylperoxy)cyclohexane, tert-butyl peroxyacetate, cumylperoxyneodecanoate, tert-amyl peroxyneodecanoate, tert-amylperoxypivalate, tert-butyl peroxyneodecanoate, tert-butyl permaleate,tert-butyl peroxypivalate, tert-butyl peroxyisononanoate,diisopropylbenzene hydroperoxide, cumene hydroperoxide, tert-butylperoxybenzoate, methyl isobutyl ketone hydroperoxide,3,6,9-triethyl-3,6,9-trimethyl-triperoxocyclononane and2,2-di(tert-butylperoxy)butane. Azoalkanes (diazenes), azodicarboxylicesters, azodicarboxylic dinitriles such as azobisisobutyronitrile andhydrocarbons which decompose into free radicals and are also referred asC-C initiators, e.g. 1,2-diphenyl-1,2-dimethylethane derivatives and1,1,2,2-tetramethylethane derivatives, are also suitable. It is possibleto use either individual initiators or mixtures of various initiators. Alarge range of initiators, such as peroxides, are commerciallyavailable, for example the products of Akzo Nobel offered under thetrade names Trigonox® or Perkadox®.

Peroxidic polymerization initiators for use in the present technologyinclude, for example, 1,1-di(tert-butylperoxy)cyclohexane,2,2-di(tert-butylperoxy)butane, tert-butylperoxy-3,5,5-trimethylhexanoate, tert-butyl peroxybenzoate,2,5-dimethyl-2,5-di(tert-butylperoxy)hexane, tert-butyl cumyl peroxide,di-tert-butyl peroxide and2,5-dimethyl-2,5-di(tert-butylperoxy)hex-3-yne, and particularpreference is given to using tert-butyl peroxy-3,5,5-trimethylhexanoate,di-(2-ethylhexyl)peroxydicarbonate or tert-butylperoxy-2-ethylhexanoate.

The initiators can be employed individually or as a mixture inconcentrations of 0.1 mol/t to 50 mol/t of polyethylene produced,including rom 0.2 mol/t to 20 mol/t, in each reaction zone. In someembodiments the free-radical polymerization initiator, which is fed to areaction zone, is a mixture of at least two different azo compounds ororganic peroxides. If such initiator mixtures are used they may be fedto all reaction zones. There is no limit to the number of differentinitiators in such a mixture. In some embodiments the mixtures arecomposed of from two to six different initiators, such as two, three orfour different initiators. In certain embodiments, the mixtures maycomprises initiators which have different decomposition temperatures.

It can be beneficial in some embodiments to use the initiators in adissolved state. Examples of suitable solvents are ketones and aliphatichydrocarbons, in particular octane, decane and isododecane and alsoother saturated C₈-C₂₅-hydrocarbons. The solutions comprise theinitiators or initiator mixtures in proportions of 2 to 65% by weight,including 5 to 40% by weight and 8 to 30% by weight.

In one embodiment of the present disclosure, the polymerization iscarried out in the presence of a chain transfer agent. Chain transferagents are chemical compounds which may interact with a growing polymerchain, terminate the further growth of the growing polymer chain andinduce the growth of another growing chain. Accordingly, the molecularweight of the polymers to be prepared can be altered by the addition ofchain transfer agents which are also called modifiers or regulators.Examples of suitable chain transfer agents are hydrogen, aliphatic andolefinic hydrocarbons, e.g. propane, butane, pentane, hexane,cyclohexane, propene, 1-butene, 1-pentene or 1-hexene, ketones such asacetone, methyl ethyl ketone (2-butanone), methyl isobutyl ketone,methyl isoamyl ketone, diethyl ketone or diamyl ketone, aldehydes suchas formaldehyde, acetaldehyde or propionaldehyde and saturated aliphaticalcohols such as methanol, ethanol, propanol, isopropanol or butanol ormixtures of these compounds. The amount of chain transfer agent fed tothe tubular reactor is, in certain embodiments, from 0.2 kg/t preparedpolymer to 40 kg/t prepared polymer, such as 0.4 kg/t prepared polymerto 20 kg/t prepared polymer. In one embodiment a chain transfer agentcomprising at least one aldehyde or at least one ketone is employed andthe amount of added aldehydes and ketones is from 0.4 kg/t of preparedpolymer to 10 kg/t of prepared polymer.

In a further embodiment of the present disclosure, propionic aldehyde isused as sole chain transfer agent or the chain transfer agent is amixture of propionic aldehyde and one or more olefinic hydrocarbons.

The high-pressure polymerization is carried out at pressures in a rangefrom 180 MPa to 340 MPa, for instance at pressures of 200 MPa to 320MPa. The polymerization temperatures are, in some embodiments, in arange from 100° C. to 350° C., including from 120° C. to 340° C. andfrom 150° C. to 330° C.

The compression of the reaction gas composition to the polymerizationpressure is carried out by a combination of a primary compressor and asecondary compressor in which the primary compressor, in certainembodiments, first compresses the reaction gas composition to a pressureof 10 MPa to 50 MPa and the secondary compressor further compresses thereaction gas composition to the polymerization pressure of 180 MPa to340 MPa. In some embodiments, the primary compressor and the secondarycompressor are multistage compressors. It is further possible toseparate one or more stages of one or both of the compressors and dividethe stages into separated compressors. However, a series of one primarycompressor and one secondary compressor may be used for compressing thereaction gas composition to the polymerization pressure. In such cases,the whole primary compressor may be designated as the primarycompressor. However, one or more first stages of the primary compressor,which compress the recycle gas from the low-pressure product separatorto the pressure of the fresh ethylene feed, may be designated as thebooster compressor and the one or more subsequent stages may bedesignated as the primary compressor, although the booster compressorand the subsequent stages may comprise a single apparatus. A secondarycompressor may be referred to as a hyper compressor. In someembodiments, the capacity of the secondary compressor, defined as thefeed rate of compressed reaction gas composition from the compressorcombination to the tubular reactor, is from 80 t/h to 210 t/h, includingfrom 100 t/h to 180 t/h and from 120 t/h to 160 t/h.

In one embodiment of the present disclosure, the entire reaction gascomposition provided by the secondary compressor is fed to the inlet ofthe tubular reactor. In another embodiment, only a part of the reactiongas composition compressed by the secondary compressor is fed to theinlet of the tubular reactor and the remainder of the reaction gascomposition compressed by the secondary compressor is fed as one or moreside streams to the tubular reactor downstream of the inlet of thetubular reactor. In such a set-up, 30 to 90% by weight, including 40 to70% by weight, of the reaction gas composition provided by the secondarycompressor is fed to the inlet of the tubular reactor and 10 to 70% byweight, such as 30 to 60% by weight, of the reaction gas compositionprovided by the secondary compressor is fed as one or more side streamsto the tubular reactor downstream of the inlet of the tubular reactor.

The process of the present disclosure may be conducted in a tubularreactor comprising a specific reactor surface area A_(sp) of the tubularreactor, e.g. the ratio of the area of the inner surface of the tubularpolymerization reactor to the feed rate of reaction gas composition fedfrom the compressor combination to the tubular reactor, is 2 m²/(t/h) to5.5 m²/(t/h), such as 3 m²/(t/h) to 5 m²/(t/h), and 3.5 m²/(t/h) to 4.5m²/(t/h). In certain embodiments where the polymerization is carried outin a manner such that a part of the reaction gas composition is fed asone or more side streams to a downstream part of the tubular reactor,the feed rate for calculating the specific reactor surface area A_(sp)is the combination of the feed rate of the reaction gas composition fedto the inlet of the tubular reactor and the feed rates of the sidestreams.

Tubular reactors for use in the present technology may be long,thick-walled pipes from about 0.5 km to 5 km in length, such as from 1km to 4 km and from 1.5 km to 3 km. The inner diameter d_(i) of thetubular reactor may be from 50 mm to 120 mm, including from 60 mm to 100mm. In case the tubular reactor is designed to be operated with one ormore reaction gas side streams, the farthest downstream part of thetubular reactor, which is passed by the total reaction mixture, maycomprise an inner diameter d_(i) in the range from 50 mm to 120 mm, suchas from 60 mm to 100 mm. Suitable tubular reactors may have alength-to-diameter ratio of greater than 1000, including from 10000 to50000 and from 20000 to 35000. In certain embodiments, the tubularreactor is composed of tubes of a length from 5 m to 25 m, such as from10 m to 22 m and from 15 m to 20 m. The individual tubes of the tubularreactor may be flanged together. The tube can also be flanged to a bend,for instance to a 180° bend. Such 180° bends may have a relatively smallradius, i.e. a ratio R/d_(o) of 4 or less, with R being the radius ofcurvature of the bending and d_(o) being the outer diameter of the tube,for the purpose of saving space.

The tubular reactor employed in the process of the present disclosure ischaracterized in that the specific ratio RD_(sp), which is the ratio ofthe outer diameter d_(o) to the inner diameter d_(i) divided by thedesign pressure p_(des) according to equation

${{RD}_{sp} = \frac{d_{o}}{d_{i}*p_{des}}},$

is in the range from 0.0050 MPa⁻¹ to 0.0069 MPa⁻¹. In certainembodiments, the specific ratio RD_(sp) is in the range from 0.0053MPa⁻¹ to 0.0066 MPa⁻¹, such as from 0.0056 MPa⁻¹ to 0.0063 MPa⁻¹. Incase the tubular reactor is operated in a manner that the total amountof reaction gas composition is fed to the inlet of the tubular reactor,in some embodiments the tubular reactor has the same inner and outerdiameters over the total length of the tubular reactor, although in theregion of a bend the outer diameter may be larger than that of thestraight tubes for compensating for an uneven wall thicknessdistribution caused by the bending process. In case the polymerizationis carried out in a manner that only a part of the reaction gascomposition is fed to the inlet of the tubular reactor and the remainderof the reaction gas composition is fed as one or more side streams to adownstream part of the tubular reactor, it is possible that the tubularreactor has the same inner and outer diameters over the total length ofthe tubular reactor. For operating the tubular reactor with reaction gasside streams, in certain embodiments the inner diameter d_(i) increasesdown-stream of the feeding points of the side streams. In such a casethe process of the present disclosure may be carried out in a tubularreactor of which at least the most downstream part of the reactor has aspecific ratio RD_(sp) in the range from 0.0050 MPa⁻¹ to 0.0069 MPa⁻¹,including from 0.0053 MPa⁻¹ to 0.0066 MPa⁻¹ and from 0.0056 MPa⁻¹ to0.0063 MPa⁻¹, where all parts of the reactor have a specific ratioRD_(sp) in the range from 0.0050 MPa⁻¹ to 0.0069 MPa⁻¹, such as from0.0053 MPa⁻¹ to 0.0066 MPa⁻¹ and from 0.0056 MPa⁻¹ to 0.0063 MPa⁻¹,although the inner and outer diameters may increase from the inlet tothe outlet of the tubular reactor.

The specific ratio RD_(sp) is a measure for the ratio of the outerdiameter d_(o) to the inner diameter d_(i), referenced to the designpressure p_(des), such that a decreasing ratio results in a thinning ofthe reactor walls for a given design pressure p_(des). Conversely,because higher design pressures require higher wall strengths,constructing a tubular reactor for a higher design pressure p_(des) withan identical specific ratio RD_(sp) results in the construction of atubular reactor with an increasing ratio d_(o) to d_(i) or, if the innerdiameter d_(i) is kept constant, with a larger outer diameter d_(o). Incertain embodiments, designing a tubular reactor with a specific ratioRD_(sp) as defined in the present disclosure results in relatively thinreactor walls and accordingly facilitates the heat transfer through thereactor wall.

The design pressure p_(des) of a tubular reactor is the maximum pressurefor which the reactor is designed for continuous usage. Tubular reactorsfor carrying out the process of the present disclosure have, in certainembodiments, a design pressure of 240 MPa to 400 MPa, including 260 MPato 380 MPa and 280 MPa to 360 MPa.

For carrying out the polymerization according to the present disclosure,steel of a suitable composition and thermal profile should be utilizedto carry out an adequate heat-treatment of the pre-formed bars fromwhich the tubes are produced. Iron-based alloys and methods forpreparing tube components from these alloys are disclosed in EP 1 529853 A2. In one embodiment of the present disclosure, the final steel ofwhich the tubular reactor consist has an tensile strength R_(m), i.e. anultimate tensile strength of at least 1100 N/mm², including at least1150 N/mm² and at least 1200 N/mm², as determined according to EN ISO6892-1:2009. In further embodiments, the steel has a 0.2% tensile yieldstrength of at least 1000 N/mm², such as at least 1100 N/mm² determinedaccording to EN ISO 6892-1:2009, and a 0.2% tensile yield strength at300° C. of at least 900 N/mm², more preferably at least 940 N/mm²,determined according to EN ISO 6892-2:2011.

In certain embodiments, the tubular reactor employed in the presentdisclosure has an inner surface roughness (Ra) of 2 μm or less, such as1 μm or less and 0.4 μm or less, determined according to DIN EN ISO4287:2010. The low roughness may be achieved by grinding, polishing,lapping and/or honing of surface.

Conducting a polymerization or copolymerization of ethylenicallyunsaturated monomers according to the process of the present disclosureadvantageously results in a higher conversion of monomers per pass ofthe reactor than carrying out such a polymerization under identicalconditions in a tubular reactor of the same dimensions, but which doesnot fulfill the required conditions for the specific ratio RD_(sp) andthe roughness Ra of the inner surface. By increasing the conversion ofmonomers per pass of the reactor, the polymer output of the reactorrises. In some embodiments of the present disclosure, the polymer outputis 2 to 5% higher than when carrying out such a polymerization underidentical conditions in a tubular reactor of the same dimensions havinga surface roughness (Ra) of more than 2 μm and a specific ratio RD_(sp)of more than 0.0069 MPa⁻¹.

Commonly the polymerization apparatus comprises, besides thepolymerization reactor and the combination of compressors, ahigh-pressure gas recycle line and a low-pressure gas recycle line forrecycling unreacted monomers to the polymerization process. The reactionmixture obtained in the polymerization reactor is transferred to a firstseparation vessel, frequently called a high-pressure product separator,and separated into a gaseous fraction and a liquid fraction at apressure of 15 MPa to 50 MPa. The gaseous fraction withdrawn from thefirst separation vessel is fed via the high-pressure gas recycle line tothe suction side of the secondary compressor. In the high-pressure gasrecycle line, the gas may be purified by several purifications steps forremoving undesired components such as entrained polymer or oligomers.The liquid fraction withdrawn from the first separation vessel, whichmay comprise dissolved monomers such as ethylene and comonomers in anamount of 20 to 40% by weight, is transferred to a second separationvessel, frequently called low-pressure product separator, and furtherseparated, at reduced pressure, at an absolute pressure in a range from0.1 to 0.5 MPa, with respect to the polymeric and gaseous components.The gaseous fraction may be withdrawn from the second separation vesseland fed via a low-pressure gas recycle line to the primary compressor.The low-pressure gas recycle line may comprise several purificationssteps for purifying the gas from the undesired components.

The recycled gas coming from the low-pressure gas recycle line may becompressed in the first stages of the primary compressor to the pressureof the fresh feed of ethylenically unsaturated monomers, such asethylene, and combined with the fresh gas feed. In some embodiments, thecombined gases are further compressed in the primary compressor to apressure of 10 MPa to 50 MPa. In certain embodiments, the primarycompressor comprises five or six compression stages, two or three beforeadding the fresh gas and two or three after adding the fresh gas. Thesecondary compressor may comprise two stages; a first stage, whichcompresses the gas from about 30 MPa to about 120 MPa, and a secondstage, which further compresses the gas from about 120 MPa to the finalpolymerization pressure.

Tubular reactors for use in the present technology may have at least tworeaction zones, such as from 2 to 6 reaction zones and from 2 to 5reaction zones. The number of reaction zones is given by the number offeeding points for the initiators or initiator mixtures. Such a feedingpoint can, for example, be an injection point for a solution of azocompounds or organic peroxides. The initiator is added to the reactorand decomposes into free radicals, which initiate polymerization. Theheat of the reaction increases the reaction mixture temperature andincreases the rate of decomposition of the free-radical initiators andaccelerates polymerization until essentially all of the free-radicalinitiator is consumed. The temperature then decreases since thetemperature of the reactor walls is lower than that of the reactionmixture. Accordingly, the part of the tubular reactor downstream of aninitiator feeding point in which the temperature rises is a reactionzone, while the part thereafter, in which the temperature decreasesagain, is predominantly a cooling zone. The amount and nature of addedfree-radical initiators determines how much the temperature rises. Insome embodiments, the temperature rise is set to be in the range from70° C. to 170° C. in the first reaction zone and 50° C. to 130° C. forthe subsequent reaction zones, depending on the product specificationsand the reactor configuration.

FIG. 1 shows schematically a set-up for a polymerization reactor inwhich the entire reaction gas composition compressed by the secondarycompressor is fed as one stream to the inlet of the tubular reactor.

The fresh ethylene, which may be under a pressure of 1.7 MPa, isinitially compressed to a pressure of about 30 MPa by means of a primarycompressor (1 a) and then compressed to the reaction pressure of about300 MPa using a secondary compressor (1 b). The chain transfer agent(CTA) is added to primary compressor (1 a). The reaction mixture leavingthe primary compressor (1 b) is fed to pre-heater (2), where thereaction mixture is preheated to the reaction start temperature of fromabout 120° C. to 220° C., and then conveyed to the inlet (3) of thetubular reactor (4).

The tubular reactor (4) is a long, thick-walled pipe with coolingjackets to remove the liberated heat of reaction from the reactionmixture by means of a coolant circuit (not shown).

The tubular reactor (4) shown in FIG. 1 has four spatially separatedinitiator injection points (5 a), (5 b), (5 c) and (5 d) for feedinginitiators or initiator mixtures PX1, PX2, PX3 and PX4 to the reactorcomprising four reaction zones. By feeding suitable free-radicalinitiators, which decompose at the temperature of the reaction mixture,to the tubular reactor the polymerization reaction starts.

The reaction mixture leaves the tubular reactor (4) through ahigh-pressure let-down valve (6) and passes through a post reactorcooler (7). Thereafter, the resulting polymer is separated off fromunreacted ethylene and other low molecular weight compounds (monomers,oligomers, polymers, additives, solvent, etc.) by means of a firstseparation vessel (8) and a second separation vessel (9), discharged andpelletized via an extruder and granulator (10).

The ethylene and comonomers which have been separated off in the firstseparation vessel (8) are fed back to the low-pressure side of thetubular reactor (4) in the high-pressure circuit (11) at 30 MPa. In thehigh-pressure circuit (11), the gaseous material separated from thereaction mixture is first freed from other constituents in at least onepurification stage and then added to the monomer stream between primarycompressor (1 a) and secondary compressor (1 b). FIG. 1 shows onepurification stage consisting of a heat exchanger (12) and a separator(13). In some embodiments, a plurality of purification stages may beused. The high-pressure circuit (11) may be used to separate waxes.

The ethylene separated off in the second separation vessel (9) furthercomprises, inter alia, the majority of the low molecular weight productsof the polymerization (oligomers) and the solvent, may be worked up inthe low-pressure circuit (14) at an absolute pressure of about 0.1 to0.5 MPa in a plurality of separators, with a heat exchanger beinglocated between each of the separators. FIG. 1 shows two purificationstages consisting of heat exchangers (15) and (17) and separators (16)and (18). It is possible to use only one purification stage or,alternatively, more than two purification stages. The low-pressurecircuit (14) may be used to separate oil and waxes.

FIG. 2 shows schematically a set-up for a polymerization reactor inwhich a part of the reaction gas composition compressed by the secondarycompressor is fed to the inlet of the tubular reactor and a part is fedas side stream to the tubular reactor down-stream of the inlet of thetubular reactor without however restricting the invention to theembodiments described therein.

The set-up shown in FIG. 2 is a modification of the set-up shown in FIG.1 in which a part of the reaction mixture leaving the primary compressor(1 b) is not fed to preheater (2) but is branched off and conveyed to afurther heat-exchanger (20). In the configuration shown in FIG. 2, thereaction mixture passing the heat-exchanger (20) is divided in 3 partialstreams which are fed via lines (21 a), (21 b) and (21 c) as sidestreams into the tubular reactor (4) at side stream injection points (22a), (22 b) and (22 c), which are located a short distance upstream ofinitiator injection points (5 b), (5 c) and (5 d). In some embodiments,the side streams fed via lines (21 a), (21 b) and (21 c) into tubularreactor (4) are utilized to cool the reaction mixture within the tubularreactor (4). Accordingly, heat-exchanger (20) normally operates as acooler and cools the reaction mixture leaving heat-exchanger (20) to atemperature in a range from 20° C. to 60° C.

Different configurations for a tubular polymerization reactor inaccordance with the present technology are possible, including the useoxygen or air as an initiator instead of peroxidic initiators. For suchpolymerizations, oxygen or air may be added to the reaction gascomposition in the primary compressor.

The polymers obtained by the process according to the present disclosureare, in some embodiments, ethylene homopolymers or copolymers having agood balance of mechanical and optical properties for use in thin films.

In certain embodiments, the polymers obtained by the process of thepresent disclosure may be converted to shaped articles such as films,laminating films or sheets, fibers, cables or wires or molded parts.Appropriate methods for the preparation of the shaped articles includeextrusion molding, extrusion coating, blow molding, rotomolding and/orinjection molding. The polymers of the present technology are suitablefor producing films, e.g. on blown film machines or cast film lines.Such films can be produced with good operability and a good balance ofoptical and mechanical properties. The present disclosure also describesprocesses for preparing shaped articles from the obtained polymers andprocesses in which the shaped article is a film.

By carrying out the high-pressure polymerization in a tubular reactorwhich has a specific ratio RD_(sp) of from 0.0050 MPa⁻¹ to 0.0069 MPa⁻¹and a relatively small wall thickness, and a smooth inner surface of asurface roughness Ra of 2 μm or less, it is, when operating in a tubularreactor of a suitable specific reactor surface area A_(sp) of 2 m²/(t/h)to 5.5 m²/(t/h), not only possible to prepare ethylene polymers with ahigh conversion rate but the process is suitable for preparing LDPE filmgrades of good operability and a good balance of optical and mechanicalproperties with a high conversion rate.

EXAMPLES

The roughness Ra of the inner surface of the reactor tubes wasdetermined according to DIN EN ISO 4287:2010.

The tensile strength Rm of the steel for constructing the tubularreactor was determined according to EN ISO 6892-1:2009.

The 0.2% tensile yield strength at 320° C. of the steel used forconstructing the tubular reactor was determined according to EN ISO6892-2:2011.

Density was determined according to DIN EN ISO 1183-1:2004, Method A(Immersion) with compression molded plaques of 2 mm thickness. Thecompression molded plaques were prepared with a defined thermal history:pressed at 180° C., 20 MPa for 8 min with crystallization in boilingwater for 30 min.

The melt flow rate (MFR) was determined according to DIN EN ISO1133:2005, condition D at a temperature of 190° C. under a load of 2.16kg.

Haze was determined according to ASTM D 1003-00 using 50 μm thicknessblown film extruded at a melt temperature of 180° C. and a blow-up ratioof 2:1.

Gloss was determined at 60° C. according to ASTM D 2457-03 using 50 μmthickness blown film extruded at a melt temperature of 180° C. and ablow-up ratio of 2:1.

The gel count was determined by preparing a cast film, analyzing thefilm defects by means of an optical scanning device and classifying andcounting the film defects according to their size (circle diameter). Thefilms were prepared by an extruder (type ME20) equipped with a chillroll and winder, model CR-9, and analyzed by an optical film surfaceanalyzer with flash camera system, model FTA100 (all components producedby OCS Optical Control Systems GmbH, Witten, Germany). The apparatus hadthe following characteristics:

-   -   screw diameter: 20 mm;    -   screw length: 25 D;    -   compression ratio: 3:1;    -   screw layout 25 D: 10 D feeding, 3 D compression, 12 D metering;    -   dimensions: 1360×650×1778 mm³ (L×W×H; without die);    -   die width (slit die): 150 mm;    -   resolution: 26 μm x 26 μm;        and was operated under the following conditions    -   T 1: 230° C.;    -   T 2 : 230° C.;    -   T 3 : 230° C.;    -   T 4 : (adapter) 230° C.;    -   T 5: (die) 230° C.;    -   die slit die: 150 mm;    -   take off speed: 3.0 m/min;    -   screw speed: to be adjusted to film thickness (50 μm);    -   throughput:1.0 to 1.5 kg/h (target 1.15 kg/h);    -   air shower on: −5 m³/h,    -   chill roll temperature: 50° C.;    -   vab chill roll: 4 N;    -   winding tensile force: 4 N,    -   draw off strength : 5 N;    -   camera threshold threshold 1: 75%-threshold 2: 65%.

For starting the measurement, the extruder and take off unit were set tothe specified conditions and started with a material having a known gellevel. The film inspection software was started when the extruder showedsteady conditions of temperature and melt pressure. After havingoperated the extruder with the starting material for at least half anhour or after the gel count having reached the known gel level, thefirst sample to measure was fed to the extruder. After having reached astable gel level for 45 minutes, the counting process was started untilthe camera had inspected an area of at least 3 m² of film. Thereafter,the next sample was fed to the extruder and after having reached astable gel count for 45 minutes the counting process for the next samplewas started. The counting process was set for all samples in a way thatthe camera inspected an area of at least 3 m² of film and the number ofmeasured defects per size-class was normalized to 1 m² of film.

The draw down thickness was determined by preparing a blown film withslowly increasing take-off speed until the film had broken. The filmthickness at “break point” is reported as draw down thickness. The filmswere prepared by a film blowing line having the followingcharacteristics:

-   -   single screw extruder with grooved feed section: 50mm×30 D;    -   barrier screw with mixing elements at the screw tip;    -   spiral mandrel die: 120 mm×1 mm;    -   extruder temperature: MFR 3-0.7: 180° C.;    -   throughput: 35 kg/hr;    -   film blow up ratio: 2.5;    -   single lip cooling ring (suitable for blow up ratios from 2:1 to        4:1);    -   no internal bubble cooling system;    -   height adjustable calibrating basket with Teflon® rolls;    -   lay flat unit equipped with rolls (CFRP guide rolls, V-shape        side positioning guidance by CFRP rolls; CFRP=Carbon Fiber        Reinforced Plastic).

For determining the draw down thickness, the film preparation wasstarted with a take-off speed of 3 m/min. Without changing processingconditions such as frost line, blow up ratio, and film width, thetake-off speed was increased by a rate of 2 m/min every 20 seconds untilthe film broke or showed the formation of holes. The draw down thicknesswas measured at a distance of 70 cm prior to the location of the filmbreak or hole formation. Measurements were made circumferentially indistances of 3 to 4 cm and the average of these measurements wasreported as the draw down thickness.

Example 1

A LDPE film grade was produced by continuous polymerization of ethyleneas ethylenically unsaturated monomer in a high-pressure tubular reactorof the design shown in FIG. 1. The reactor had a total length of 2140 mand a design pressure of 300 MPa. The tubular reactor was composed ofreactor tubes having a length of 17 m and an inner surface roughness ofRa=0.1 μm. The steel of the tubes had a tensile strength R_(m) of 1210N/mm² and a 0.2% tensile yield strength at 320° C. of 920 N/mm², thetubular reactor had an outer diameter d_(o) of 137 mm and an innerdiameter d_(i) of 75 mm, resulting in a specific ratio RD_(sp) of 0.0061MPa⁻¹. The peroxidic polymerization initiators were metered into thetubular reactor at four positions using isododecane as additionaldiluent. Propionaldehyde was added as a chain transfer agent to thefresh ethylene stream entering the primary compressor (1 a) in an amountof 1.4 kg per ton (t) of produced polyethylene. The secondary compressor(1 b) was operated with a throughput of 125 t reaction gas compositionper hour. Accordingly, the polymerization was carried out with aspecific reactor surface area of 4.03 m²/(t/h).

The compressed reaction gas composition was heated to 160° C. inpre-heater (2) and fed to the inlet of the tubular reactor having apressure of 260 MPa. For removing the generated heat of polymerization,cooling water was circulated through the cooling jackets attached to theouter surface of the reactor tubes (not shown in FIG. 1). The reactionmixture discharged from the tubular reactor (4) was passed through apost reactor cooler (7) and separated from volatiles in two steps via afirst separation vessel (8) and a second separation vessel (9). Thedegassed liquid polymer was conveyed to an extruder and granulator (10)to form LDPE pellets. Within a period of 24 hours of continuousproduction, 975 t of LDPE were obtained having the properties summarizedin Table 1. Accordingly, the conversion rate of ethylene per passthrough the reactor was 32.5%.

Comparative Example A

A film grade LDPE of the same density and melt flow rate as the filmgrade LDPE prepared in Example 1 was produced by continuouspolymerization of ethylene as ethylenically unsaturated monomer in ahigh-pressure tubular reactor of the design shown in FIG. 1 having intotal a length of 1125 m and a design pressure of 340 MPa. The tubularreactor was composed of reactor tubes having a length of 10 m and aninner surface roughness of Ra=3.8 μm. The steel of the tubes had atensile strength R_(m) of 990 N/mm² and a 0.2% tensile yield strength at320° C. of 720 N/mm², allowing to construct the tubular with an outerdiameter d_(o) of 96 mm and an inner diameter d_(i) of 40 mm, thusresulting in a specific ratio RD_(sp) of 0.0071 MPa⁻¹. The peroxidicpolymerization initiators were metered into the tubular reactor at fourpositions using isododecane as additional diluent. Propionaldehyde wasadded as a chain transfer agent to the fresh ethylene stream enteringthe primary compressor (1 a) in an amount of 1.4 kg per t of producedpolyethylene. The secondary compressor (1 b) was operated with athroughput of 32 t reaction gas composition per hour. Accordingly, thepolymerization was carried out with a specific reactor surface area of4.42 m²/(t/h).

The polymerizations were carried out as described in Example 1. Within aperiod of 24 hours of continuous production, 240 t of LDPE were obtainedhaving the properties summarized in Table 1. Accordingly, the conversionrate of ethylene per pass through the reactor was 31.3%.

Comparative Example B

The polymerization of Comparative Example A was repeated except that amixture of propionaldehyde and propene in a weight ratio of 1:10 wasadded as a chain transfer agent to the fresh ethylene stream enteringthe primary compressor (1 a) in a total amount of 8 kg per t of producedpolyethylene. Within a period of 24 hours of continuous production, 210t of LDPE were obtained having the properties summarized in Table 1.Accordingly, the conversion rate of ethylene per pass through thereactor was 27.3%.

TABLE 1 Comparative Comparative. Example 1 Example A Example B Density[g/cm³] 0.9232 0.9231 0.9230 MFR_(2.16) [g/10 min] 0.75 0.76 0.75 Haze[%] 6.5 8.3 6.7 Gloss 97 69 96 Draw down thickness [μm] 17 30 18 Gelcount <200 μm [1/100 m²] 365 542 340 200-400 μm [1/100 m²] 75 119 77400-800 μm [1/100 m²] 6 11 7

The comparison of Example 1 and Comparative Examples A and Bdemonstrates that by reducing the specific ratio RD_(sp) of the tubularreactor in combination with reducing the roughness of the inner surfaceit is possible to prepare LDPE with a higher conversion of ethylene perpass of the reactor without losses in the polymer properties, while thecomparison of Comparative Examples A and B shows that conversion rateper pass of the reactor and polymer properties cannot be variedindependently for a LPDE grade of defined density and melt flow rate.

What is claimed is:
 1. A tubular reactor for polymerizing orcopolymerizing ethylenically unsaturated monomers, comprising: a lengthranging from 0.5 km to 5 km, an inner diameter d, ranging from 50 mm to120 mm, and a length-to-diameter ratio ranging from 1000 to 50000; aninlet and at least two spatially separated initiator injection pointsdownstream of the inlet; and at least a portion of the length having aspecific ratio RD_(sp) ranging from 0.0050-0.0069 MPa⁻¹, RD_(sp) beingthe ratio of an outer diameter d_(o) to an inner diameter d_(i) dividedby a design pressure p_(des) according to the equation:${RD}_{sp} = {\frac{d_{o}}{d_{i}*p_{des}}.}$
 2. (canceled)
 3. (canceled)4. (canceled)
 5. (canceled)
 6. The tubular reactor of claim 1, whereinthe design pressure p_(des) of the tubular reactor ranges from 240 MPato 400 MPa.
 7. The tubular reactor of claim 1, wherein the pipe has aratio of an outer diameter d_(o) to the inner diameter d_(i) rangingfrom 1.2 to 2.76.
 8. The tubular reactor of claim 1, further comprisingan inner surface which has a surface roughness Ra of 2 μm or less,determined according to DIN EN ISO 4287:2010.
 9. The tubular reactor ofclaim 1, having a design feed rate of reaction gas composition to thetubular reactor ranges from 80 t/h to 210 t/h.
 10. The tubular reactorof claim 1, having a design feed rate of reaction gas composition to thetubular reactor ranges from 100 t/h to 180 t/h.
 11. The tubular reactorof claim 1, having a design feed rate of reaction gas composition to thetubular reactor ranges from 120 t/h to 160 t/h.
 12. The tubular reactorof claim 1, further comprising a tensile strength R_(m) of at least 1100N/mm² determined according to EN ISO 6892-1:2009.
 13. (canceled)
 14. Thetubular reactor of claim 1, having an outer diameter d_(o), an innerdiameter d_(i), or an outer diameter d_(o) and an inner diameter d_(i)that varies over the length of the pipe.
 15. A process for polymerizingor copolymerizing ethylenically unsaturated monomers, comprising:feeding a reaction gas into a tubular reactor having a specific ratioRD_(sp) ranging from 0.0050-0.0069 MPa⁻¹, RD_(sp) being the ratio of theouter diameter d_(o) to the inner diameter d_(i) divided by the designpressure p_(des) according to the equation:${{RD}_{sp} = \frac{d_{o}}{d_{i}*p_{des}}};$ feeding an initiator intothe tubular reactor in a concentration ranging from 0.1 mol/t to 50mol/t; and polymerizing the reaction gas in the tubular reactor at atemperature ranging from 100° C. to 350° C. and a pressure ranging from180 MPa to 340 MPa.
 16. (canceled)
 17. (canceled)
 18. (canceled)
 19. Theprocess of claim 15, further comprising: feeding a 30-90% by weightportion of the reaction gas to an inlet of the tubular reactor; andfeeding a 10-70% by weight portion of the reaction gas to one or moreside streams of the tubular reactor downstream of the inlet of thetubular reactor.
 20. The process of claim 19, wherein the tubularreactor has an inner surface which has a surface roughness Ra of 2 μm orless, determined according to DIN EN ISO 4287:2010.
 21. The tubularreactor of claim 1, comprising at least a portion of the length having aspecific ratio RD_(sp) ranging from 0.0053 MPa⁻¹ to 0.0066 MPa⁻¹. 22.The tubular reactor of claim 1, comprising at least a portion of thelength having a specific ratio RD_(sp) ranging from 0.0056 MPa⁻¹ to0.0063 MPa⁻¹.
 23. The process of claim 15, comprising at least a portionof the tubular reactor has a specific ratio RD_(sp) ranging from 0.0053MPa⁻¹ to 0.0066 MPa⁻¹.
 24. The process of claim 15, comprising at leasta portion of the tubular reactor has a specific ratio RD_(sp) rangingfrom 0.0056 MPa⁻¹ to 0.0063 MPa⁻¹.